Method for manufacturing propylene and low-sulfur fuel oil components

Abstract

A method for producing propylene and low-sulfur fuel oil components comprises the steps of: contacting a heavy feedstock oil with a solvent for extractive separation to obtain deasphalted oil and de-asphalt; contacting the deasphalted oil and optionally a light feedstock oil with a catalytic conversion catalyst for reaction in the absence of hydrogen to obtain a reaction product containing propylene; separating the reaction product to obtain catalytic cracking distillate, and hydrodesulfurizing the catalytic cracking distillate to obtain a low-sulfur hydrodistillate, wherein the low-sulfur hydrodistillate and / or de-asphalt is suitable for use as a fuel oil component. This process enables the conversion of saturated hydrocarbons in the heavy feedstock into propylene and eliminates the use of saturated hydrocarbons in the fuel oil component, thereby providing better economic and social advantages.

Description

Technology Field

[0001] Cross-reference of related applications

[0002] The present invention claims priority to patent application No. 201911014995.6 filed on October 24, 2019, titled “Method for further producing propylene and low-sulfur fuel oil components,” the entirety of which is incorporated by reference into the present invention.

[0003] The present invention relates to the field of catalytic conversion of hydrocarbon oils, in particular to a catalytic conversion process for converting heavy feedstock oils into light olefin and low-sulfur fuel oil components. Background Technology

[0004] With the rapid development of the social economy, the concept of environmental protection based on post-pollution treatment has been replaced by the concept of environmental protection based on pollution prevention from raw materials. Environmental pollution issues are receiving increasing attention, and relevant laws and regulations are becoming increasingly stringent. According to the International Maritime Organization (IMO) Convention on the Prevention of Pollution from Ships, starting January 1, 2020, ships worldwide must use marine fuel with a sulfur content of 0.5% or less. According to BP's forecast, global marine fuel consumption could reach approximately 300 million tons in 2020, which will pose a significant challenge to the global fuel market and major oil processing companies.

[0005] Meanwhile, the world is facing the problem of crude oil becoming increasingly heavier and inferior. It is estimated that heavy oil reserves will account for approximately 50% of global recoverable crude oil reserves after 2020. As the price gap between inferior and premium crude oil widens, how to efficiently utilize and process inferior heavy oil to produce low-sulfur marine fuels that meet higher environmental protection demands, improve yields, and meet high value-added products has become an urgent issue for refineries and suppliers, posing a challenge to traditional crude oil processing technologies.

[0006] CN102746890A discloses a method for producing marine fuel by visbreaking and fractionating feedstock oil to obtain a visbreaking blending component. This process can reduce the production cost of marine fuel. The process comprises: 1) a step of visbreaking a heavy oil component; 2) a step of fractionating the visbreaking product and collecting a fraction with a high distillation range; and 3) a step of mixing the fraction with a high distillation range with a light oil component to obtain marine fuel.

[0007] Since the price of marine fuel is lower than that of automotive diesel, it is difficult for marine fuel production to provide better economic benefits. Therefore, producing marine fuel while simultaneously generating high-value-added products such as propylene through a process that produces light olefins at high yields is of direct and significant importance to the compositional characteristics of the feedstock oil.

[0008] Therefore, in light of the global deterioration of oil quality and strict environmental protection requirements, there is a need to develop a method to produce high-value-added propylene at a high yield while producing low-sulfur marine fuel components that can meet market demand for high-quality fuel oil and enhance corporate economic benefits.

[0009] The object of the present invention is to provide a method for producing propylene and low-sulfur fuel oil components from heavy oil, which enables the conversion of saturated hydrocarbons in heavy feedstock oil into propylene and the conversion of polycyclic aromatic materials containing an aromatic core structure into fuel oil components, thereby suppressing the use of saturated hydrocarbons in fuel oil components and providing better economic and social benefits.

[0010] To achieve the above objective, the present invention provides a method for producing propylene and low-sulfur fuel oil components comprising the following steps:

[0011] (1) A step of obtaining deasphalted oil and deoiled asphalt by contacting a solvent for extraction separation with heavy feedstock oil;

[0012] (2) A step of contacting deasphalted oil and optionally light feedstock oil in a catalytic conversion reactor in the absence of hydrogen with a catalytic conversion catalyst for reaction to obtain a reaction product containing propylene;

[0013] (3) A step of separating the reaction product from step (2) to obtain catalytic cracking distillate, wherein the catalytic cracking distillate has an initial boiling point of about 200°C or higher and a final boiling point of about 550°C or lower, and a hydrogen content of about 12.0 wt% or lower; and

[0014] (4) A step of obtaining low-sulfur hydrogenated distillate oil by hydrogenating the above catalytic cracking distillate oil,

[0015] Here, low-sulfur hydrogenated distillate and / or deoiled asphalt are suitable for use as fuel oil components, and

[0016] The catalytic conversion catalyst used in step (2) comprises, based on the total weight of the catalyst, about 1-50 wt% zeolite, about 5-99 wt% inorganic oxide, and about 0-70 wt% clay, and

[0017] The reaction conditions of step (2) are a reaction temperature of about 460-750°C, preferably about 480-700°C, and about 10-100 h -1 , preferably about 30-100 h -1 It includes a space velocity per unit weight of, or a reaction time of about 1-10 seconds, preferably about 2-8 seconds; and a weight ratio of catalyst to oil of about 4-20, preferably about 5-12.

[0018] In the method of the present invention, asphaltenes and resins are separated from heavy feedstock oil through solvent deasphalting, and the deasphalted asphalt can be used as a fuel oil component. Since the deasphalted oil is used as a feedstock for selective catalytic cracking, catalytic cracking distillate containing single-chain polycyclic aromatic substances and propylene can be obtained to the maximum extent. The catalytic cracking distillate can be used as a fuel oil component either alone or mixed with deasphalted asphalt or heavy deasphalted oil. By using the method of the present invention, heavy feedstock oil can be converted into propylene, butylene, and fuel oil components, and the yields of dry gas and coke are significantly reduced, thereby enabling the effective utilization of petroleum resources.

[0019] In particular, when compared to the prior art, the method of the present invention can provide at least one of the following advantages:

[0020] 1. Propylene, a high value-added product, can be produced with high yield from heavy feedstock oil along with fuel oil components. By utilizing the above process, approximately 5–20 wt% of propylene and approximately 30–80 wt% of fuel oil components can be produced relative to the weight of the feedstock oil (heavy feedstock oil + light feedstock oil), thereby obtaining greater economic benefits compared to simply using feedstock oil for fuel oil blending;

[0021] 2. The production of high value-added products such as propylene may increase significantly, while the yields of dry gas and coke decrease significantly;

[0022] 3. Since the total liquid yield can be significantly increased without practically discharging slurry oil, the efficiency of petroleum resources can be improved;

[0023] 4. The catalytic conversion unit is integrated with the solvent deasphalting unit, and the characteristics of the catalytic cracking distillate oil obtained from the catalytic conversion unit can be modified by adjusting the amount of solvent used in the solvent deasphalting unit and the amount of light feedstock oil used, so the method is suitable for various amounts of heavy feedstock oil and fuel oil components. Brief explanation of the drawing

[0024] The drawings constituting part of this specification are provided to aid in understanding the invention and should not be construed as limiting. The invention may be interpreted with reference to the drawings in conjunction with the following detailed description. In the drawings: FIG. 1 shows a schematic flowchart of a preferred embodiment of the method according to the present invention. FIG. 2 shows a schematic flowchart of another preferred embodiment of the method according to the present invention. FIG. 3 shows a schematic diagram of a preferred embodiment of a solvent deasphalting unit used in the present invention. FIG. 4 shows a schematic diagram of a preferred embodiment of a catalytic conversion unit used in the present invention. Specific details for implementing the invention

[0025] The present invention will be described in more detail below with reference to specific embodiments of the invention and the accompanying drawings. It should be noted that specific embodiments of the present invention are provided merely for illustrative purposes and are not intended to limit the invention in any way.

[0026] Any specific numeric value including an endpoint of a numeric range described in the context of the present invention is not limited to that exact value, but should be interpreted to further include all values ​​close to said exact value, for example, all values ​​within ±5% of said exact value. Additionally, with respect to any numeric range described herein, any combination may be formed between endpoints of the range, between each endpoint and any specific value within the range, or between any two specific values ​​within the range to provide one or more new numeric range(s). Herein, said new numeric range(s) should also be considered as specifically described herein.

[0027] Unless otherwise noted, terms used herein have the same meaning as generally understood by those skilled in the art; and in the event that a term is defined herein and the definition differs from the general understanding of the art, the definition provided herein shall prevail.

[0028] According to the present invention, the term “catalytic cracking distillate” refers to a fraction of the product of a catalytic conversion reaction having an initial boiling point of about 200°C or higher, preferably about 250°C or higher, and a final boiling point of about 550°C or lower, preferably about 520°C or lower, most preferably about 500°C or lower, that is, a fraction having a distillation range of about 200-550°C, preferably about 250-520°C, more preferably about 250-500°C.

[0029] In the present invention, the term "fluidized bed reactor," also referred to as "fluidized reactor," should be understood in the broadest sense and encompasses all types of reactors that allow a gaseous feedstock to come into contact with solid catalyst particles for chemical reaction in a fluidized state, and includes, but is not limited to, compact beds, bubble beds, boiling beds, turbulent beds, high-velocity beds, and gas transport beds (e.g., upward-flow fluidized beds and downward-flow fluidized beds, etc.). The fluidized bed reactor may be a constant linear velocity fluidized bed reactor, an equal-diameter fluidized bed reactor, or a variable-diameter fluidized bed reactor, and may be a composite reactor comprising two or more different types of fluidized beds connected in series or parallel, such as a riser reactor or a composite reactor comprising a riser reactor combined with a compact bed. Generally, the gas velocity of the compact bed may be in the range of about 0.1–2 m / s, while the gas velocity of the riser reactor may be in the range of about 1–30 m / s (excluding catalyst).

[0030] In the context of the present invention, any or unmentioned matters other than those explicitly mentioned shall be deemed, without modification, to be identical to matters known in the art. Furthermore, any embodiment described herein may be freely combined with one or more other embodiments described herein, and any technical solution or idea obtained thereby shall be considered part of the original disclosure or the original specification of the present invention, and shall not be considered new matters not disclosed or implied herein unless it is obvious to a person skilled in the art that such combination is obviously unreasonable.

[0031] All patent and non-patent literature cited in this specification, including but not limited to textbooks and journal articles, is incorporated by reference into the present invention in its entirety.

[0032] As previously stated, the present invention provides a method for manufacturing propylene and fuel oil components comprising the following steps:

[0033] (1) A step of obtaining deasphalted oil and deoiled asphalt by contacting a solvent for extraction separation with heavy feedstock oil;

[0034] (2) A step of contacting deasphalted oil and optionally light feedstock oil in a catalytic conversion reactor in the absence of hydrogen with a catalytic conversion catalyst for reaction to obtain a reaction product containing propylene;

[0035] (3) A step of separating the reaction product from step (2) to obtain catalytic cracking distillate, and optionally liquefied gas including propylene and gasoline; and

[0036] (4) A step of obtaining low-sulfur hydrogenated distillate oil by hydrodesulfurizing the above catalytic cracking distillate oil;

[0037] Here, low-sulfur hydrogenated distillate and / or deoiled asphalt can be used as fuel oil components.

[0038] According to the present invention, step (1) is a solvent deasphalting step, which may be performed in one or two steps. In a preferred specific embodiment, step (1) is performed in two steps to obtain light deasphalted oil, heavy deasphalted oil, and de-oiled asphalt, wherein the light deasphalted oil is used as a feedstock oil for a catalytic conversion step (2), the heavy deasphalted oil may be used as a fuel oil component or for other purposes, and the de-oiled asphalt may be used as a fuel oil component or, depending on its sulfur content, as a road asphalt product.

[0039] According to the present invention, the heavy feedstock oil used in step (1) may be any heavy oil suitable for use in a catalytic conversion process, and may be selected from, for example, vacuum residue, degraded atmosphere residue, hydrogenated heavy oil, or any mixture thereof.

[0040] According to the present invention, the solvent used in step (1) may be any solvent suitable for use in the solvent deasphalting process of heavy oil, for example, selected from propane, butane, pentane, etc., or any mixture thereof, and the deasphalting solvent may be selected according to the characteristics of the catalytic cracking distillate obtained from the catalytic conversion unit.

[0041] In a preferred embodiment, the conditions of the solvent deasphalting step (1) may include a temperature of about 10-200°C, preferably about 20-180°C; a pressure of about 1.0-15.0 MPa, preferably about 2.0-10.0 MPa; and a mass ratio of solvent to heavy feedstock oil of about 1-20, preferably about 3-10 (hereinafter also referred to as the “solvent ratio”).

[0042] According to the present invention, the light feedstock oil used in step (2) may be selected from the group consisting of petroleum hydrocarbons, other mineral oils, or mixtures thereof, wherein the petroleum hydrocarbons may be selected from the group consisting of vacuum gas oil, atmospheric gas oil, coker gas oil, high-quality residue, high-quality hydrogenated heavy oil, or any mixture thereof, and the other mineral oils may be selected from the group consisting of coal liquefaction oil, tar sand oil, shale oil, or any mixture thereof, and the “high-quality residue” and “high-quality hydrogenated heavy oil” each mean hydrogenated heavy oil and residue having a hydrogen content of 11.2 weight% or more, preferably 12.0 weight% or more, most preferably 12.5 weight% or more.

[0043] According to the present invention, the catalytic conversion reactor used in step (2) may be selected from various types of fluidized bed reactors, such as a single fluidized bed reactor or a composite reactor comprising a plurality of fluidized bed reactors connected in series or in parallel. In a preferred specific embodiment, the fluidized bed reactor may be an equal-diameter riser reactor or various types of variable-diameter fluidized bed reactors, such as the reactor disclosed in Chinese Patent No. CN 1078094C.

[0044] According to the present invention, the catalytic conversion catalyst used in step (2) may comprise about 1-50 wt% of zeolite, about 5-99 wt% of inorganic oxide, and about 0-70 wt% of clay, based on the total weight of the catalyst. Preferably, the catalyst may comprise about 5-45 wt% of zeolite, and more preferably, about 10-40 wt% of zeolite, about 5-80 wt% of inorganic oxide, and about 10-70 wt% of clay.

[0045] In a preferred embodiment, the zeolite may comprise about 51-100 wt%, preferably about 70-100 wt%, of mesoporous zeolite and about 0-49 wt%, preferably about 0-30 wt%, of macroporous zeolite based on the total weight of the zeolite, and preferably the mesoporous zeolite has a silica-alumina ratio greater than about 10, preferably greater than about 50, more preferably greater than about 100. The mesoporous zeolite is preferably selected from the group consisting of ZSM-type zeolite and ZRP zeolite; the macroporous zeolite is preferably Y-type zeolite. Optionally, the zeolite may be modified with non-metallic elements such as phosphorus and / or transition metal elements such as iron, cobalt, nickel, etc. The inorganic oxide is preferably selected from the group consisting of silica, alumina, and combinations thereof; and the clay is preferably selected from kaolin and / or halloysite.

[0046] According to the present invention, step (2) is performed under effective conditions, wherein the “effective conditions” refer to conditions in which the reaction feedstock undergoes a catalytic conversion reaction comprising propylene and catalytic cracking distillate, preferably in which a reaction product comprising about 8-25% by weight of propylene and about 15-50% by weight of catalytic cracking distillate is obtained with respect to the weight of the reaction feedstock (i.e., deasphalted oil + optional light feedstock oil). In a preferred embodiment, the reaction conditions of step (2) are a reaction temperature of about 460-750°C, preferably about 480-700°C, more preferably about 480-600°C, most preferably about 500-560°C; about 5-100 h -1 , preferably about 10-70 h -1 , More preferably about 15-50 h -1 , most preferably about 18-40 h -1 It includes a space velocity per unit time of weight (e.g., for a close-bed reactor, a high-speed bed reactor, etc.), or a reaction time of about 1-10 seconds, preferably about 2-8 seconds, more preferably about 2.5-8 seconds, most preferably about 3-8 seconds (e.g., for a riser reactor); and a weight ratio of the catalytic conversion catalyst to the catalytic conversion feedstock oil of about 4-20, preferably about 5-12, preferably about 5-10, more preferably about 5-9 (also referred to as the “weight ratio of catalyst to oil” or “ratio of catalyst to oil”).

[0047] In a preferred embodiment, step (2) is performed such that the resulting reaction product has a propylene / propane mass ratio of about 4 or more, preferably about 6 or more, most preferably about 8 or more; and / or an isobutene / isobutane mass ratio of about 1 or more, preferably about 1.5 or more, most preferably about 1.8 or more.

[0048] In a preferred embodiment, step (2) is performed such that the yield of catalytic cracking distillate in the generated reaction product, i.e., the weight ratio of catalytic cracking distillate to feedstock oil for the catalytic conversion reaction (i.e., deasphalted oil and any light feedstock oil), is about 15% or more, preferably about 20% or more, more preferably about 30% or more, and about 50% or less.

[0049] As is well known to those skilled in the art, the conversion rate of feedstock oil in a catalytic conversion process is generally expressed as the sum of the yields of gas, gasoline, and coke. In the process of the present invention, the final products of the catalytic conversion process consist only of dry gas, liquefied gas, gasoline, catalytic cracking distillate, and coke. Accordingly, in this application, the conversion rate of feedstock oil is substantially equal to 100% minus the yield of catalytic cracking distillate. In turn, the conversion rate of the catalytic conversion process according to the present invention is controlled to a level of about 85% or less, preferably about 80% or less, most preferably about 70% or less, and about 50% or more.

[0050] In a particularly preferred embodiment, the process further comprises a step of separating the reaction products of step (2) from the spent catalyst, wherein the spent catalyst is recirculated to the reactor after regeneration by coke combustion and stripping, and the separated reaction products include propylene, gasoline, and catalytic cracking distillate. Methods for separating substances such as propylene from the reaction products are well known to those skilled in the art, and methods for solvent deasphalting of heavy feedstock oils are also well known to those skilled in the art and are therefore not described in detail herein.

[0051] According to the present invention, the catalytic cracking distillate obtained in step (3) has an initial boiling point of about 200°C or higher, a final boiling point of about 550°C or lower, and a hydrogen content of about 12.0 wt% or less. Preferably, the catalytic cracking distillate has an initial boiling point of about 250°C or higher, a final boiling point of about 520°C or lower, more preferably 500°C or lower, and a hydrogen content of about 11.0 wt% or less.

[0052] According to the present invention, the fuel oil component can be formed by optionally mixing hydrodesulfurized catalytic cracking distillate with de-asphalt or heavy de-asphalt oil obtained from a solvent deasphalting unit.

[0053] According to the present invention, the catalyst used in the hydrodesulfurization step (4) is preferably a catalyst comprising a Group VIB metal and / or a Group VIII metal supported on an alumina and / or amorphous silica-alumina carrier. More preferably, the catalyst used in the hydrotreatment step (4) comprises about 0-10 wt% of an additive, about 1-40 wt% of at least one Group VIII metal (calculated as a metal oxide), and about 1-50 wt% of at least one Group VIB metal (calculated as a metal oxide), with the remainder being a carrier selected from alumina and amorphous silica-alumina, wherein the additive comprises a non-metal element selected from fluorine, phosphorus, etc., a metal element selected from titanium, platinum, etc., or a combination thereof. For example, the additive may be a phosphorus-containing auxiliary agent or a fluorine-containing auxiliary agent such as ammonium fluoride. The Group VIB metal is preferably selected from molybdenum, tungsten, or a combination thereof; and the Group VIII metal is preferably selected from nickel, cobalt, or a combination thereof.

[0054] In a preferred embodiment, the conditions of the hydrodesulfurization step (4) are a reaction pressure of about 2.0-24.0 MPa, preferably about 3.0-15 MPa; a reaction temperature of about 200-500°C, preferably about 300-400°C; and about 50-5000 Nm 3 / m 3 , preferably about 200-2000 Nm 3 / m 3 Volume ratio of hydrogen to oil; approximately 0.1–30.0 h -1 , preferably about 0.2-10.0 h -1 It includes the liquid's space velocity per hour.

[0055] In a preferred embodiment, the low-sulfur hydrodistillate obtained in step (4) through hydrodesulfurization of the catalytic cracking distillate has a sulfur content of about 0.1 wt% or less, preferably about 0.05 wt% or less.

[0056] In a particularly preferred embodiment, the method according to the present invention comprises the following steps:

[0057] (1) A step of contacting heavy feedstock oil with a solvent in a single-stage solvent deasphalting unit for extraction separation to obtain deasphalted oil and deasphalted asphalt;

[0058] (2) A step of contacting deasphalted oil and optionally light feedstock oil in a catalytic conversion reactor in the absence of hydrogen with a catalytic conversion catalyst for reaction to obtain a reaction product containing propylene;

[0059] (3) a step of separating the reaction products from step (2) to obtain a liquefied gas comprising dry gas, propylene, gasoline and catalytic cracking distillate, wherein the liquefied gas may be further separated to obtain propylene, propane and C4 hydrocarbons; and

[0060] (4) A step of obtaining low-sulfur hydrogenated distillate by hydrodesulfurizing catalytic cracking distillate;

[0061] Here, the above-mentioned low-sulfur hydrogenated distillate can be used alone as a fuel oil component or blended with deoiled asphalt as a fuel oil component;

[0062] Preferably, low-sulfur hydrogenated distillate is blended with deoiled asphalt to form a fuel oil component or product.

[0063] In another particularly preferred embodiment, the method according to the present invention comprises the following steps:

[0064] (1) A step of contacting heavy feedstock oil with a solvent in a two-stage solvent deasphalting unit for extraction separation to obtain light deasphalted oil, heavy deasphalted oil and deasphalted asphalt;

[0065] (2) A step of contacting light deasphalted oil and optionally light feedstock oil in a catalytic conversion reactor in the absence of hydrogen with a catalytic conversion catalyst for reaction to obtain a reaction product containing propylene;

[0066] (3) a step of separating the reaction products from step (2) to obtain a liquefied gas comprising dry gas, propylene, gasoline and catalytic cracking distillate, wherein the liquefied gas may be further separated to obtain propylene, propane and C4 hydrocarbons; and

[0067] (4) A step of obtaining low-sulfur hydrogenated distillate by hydrodesulfurizing catalytic cracking distillate;

[0068] Here, the above-mentioned low-sulfur hydrogenated distillate may be used alone as a fuel oil component, or may be used as a fuel oil component blended with heavy deasphalted oil or de-asphalt;

[0069] Preferably, low-sulfur hydrogenated distillate is blended with heavy asphaltized oil as a fuel oil component or product.

[0070] A preferred embodiment of the method according to the present invention is described below with reference to the drawings.

[0071] In a particularly preferred embodiment, as shown in FIG. 1, heavy feedstock oil is fed through pipeline (1") to a single-stage solvent deasphalting unit (2") and contacted with a solvent (not shown) for separation to obtain deasphalted oil and deasphalted asphalt. The deasphalted asphalt is discharged through pipeline (4") and the solvent is recirculated (not shown). The deasphalted oil is optionally mixed with light feedstock oil from pipeline (5'') and then sent through pipeline (3'') to a catalytic conversion unit (6'') to contact a catalytic conversion catalyst for the reaction. After separation of the catalytic conversion products, the obtained dry gas is discharged through pipeline (7''); the obtained liquefied gas is recovered through pipeline (8") and can be further separated into propylene, propane, and C4 hydrocarbons; and the obtained gasoline is discharged through pipeline (9''); The obtained catalytic cracking distillate is recovered through a pipeline (10) and hydrodesulfurized in a hydrodesulfurization unit (11)). The hydrocatalytic cracking distillate is recovered as a fuel oil component through a pipeline (12) or blended with degreased asphalt from a pipeline (4) through a pipeline (13) or optionally further blended with light distillate from a pipeline (15) as a fuel oil component or product.

[0072] In another particularly preferred embodiment, as shown in FIG. 2, heavy feedstock oil is supplied through pipeline (1) to a two-stage solvent deasphalting unit (2) (the process method of two-stage solvent deasphalting is schematically illustrated in FIG. 3) and separated by contact with a solvent (not shown) to obtain light deasphalted oil, heavy deasphalted oil, and de-oiled asphalt. The heavy deasphalted oil and de-oiled asphalt are discharged through pipelines (14) and (4), respectively, and the solvent is recirculated (not shown). Light deasphalted oil is optionally mixed with light feedstock oil from pipeline (5) and then sent through pipeline (3) to a catalytic conversion unit (6) to come into contact with a catalytic conversion catalyst for the reaction. After the separation of the catalytic conversion products, the obtained dry gas is discharged through pipeline (7''); the obtained liquefied gas is recovered through pipeline (8'') and can be further separated into propylene, propane, and C4 hydrocarbons; the obtained gasoline is discharged through pipeline (9''); and the obtained catalytic cracking distillate is recovered through pipeline (10) and hydrodesulfurized in a hydrodesulfurization unit (11) for hydrodesulfurization treatment. The hydrocatalytic cracking distillate is recovered as a fuel oil component through pipeline (12) or blended with heavy deasphalted oil from pipeline (14) through pipeline (13) as a fuel oil component or product.

[0073] FIG. 3 illustrates a two-stage solvent deasphalting process, wherein, in the process, feedstock oil is supplied through pipeline (1') to extraction tower I of the deasphalting unit and comes into contact with a solvent from pipeline (2') for separation to obtain deasphalted asphalt containing the solvent and deasphalted oil. The deasphalted asphalt containing the solvent is sent to an asphalt evaporator through pipeline (5') for separation, the obtained deasphalted asphalt is discharged through pipeline (11'), and the solvent is discharged through pipeline (8'). The deasphalted oil is sent through pipeline (3') to extraction tower II and comes into contact with a solvent from pipeline (4') for separation to obtain light deasphalted oil and heavy deasphalted oil. The light deasphalted oil is sent through pipeline (6') to a light oil evaporator for separation, the solvent is discharged through pipeline (8'), and the light deasphalted oil is discharged through pipeline (9'); The heavy deasphalted oil is sent to a heavy oil evaporator through a pipeline (7') for separation, the solvent is discharged through a pipeline (8'), and the heavy deasphalted oil is discharged through a pipeline (10').

[0074] As shown in FIG. 4, in a particularly preferred embodiment, a pre-lifting medium is introduced through pipeline (1) into a variable diameter fluidized bed reactor (2) (e.g., the reactor disclosed in Chinese Patent No. CN 1078094C), and under the activity of the pre-lifting medium, a regenerative catalyst from a regenerative catalyst inclined pipe (16) moves upward along the reactor, and light deasphalted oil from pipeline (9') is supplied through pipeline (3) to the lower part of the first reaction zone (8) of the variable diameter fluidized bed reactor (2) together with a spray stream from pipeline (4) and mixed with the existing stream in the reactor. The feedstock oil is decomposed on the high-temperature catalyst and moves upward to the second reaction zone (9) of the variable diameter fluidized bed reactor (2) for further reaction. The generated oil gas and deactivated spent catalyst are transferred to a cyclone separator of separator (7) to achieve separation of the spent catalyst and oil gas. The oil gas passes through the main oil gas pipeline (17), and the fine powder of the catalyst is returned to the separator (7) through the dipleg of the cyclone separator. The spent catalyst in the separator (7) is transferred to the stripping section (10) and comes into contact with stripping steam from the pipeline (11). The stripped oil gas from the spent catalyst is transferred to the main oil gas pipeline (17) through the cyclone separator. The stripped spent catalyst is sent to the regenerator (13) through the spent catalyst inclined pipe (12), and main air is introduced into the regenerator through the pipeline (14) to burn the coke deposited on the spent catalyst, thereby regenerating the deactivated spent catalyst. Flue gas is discharged through the pipeline (15). The regenerated catalyst is recirculated to the variable diameter fluidized bed reactor (2) through the regenerated catalyst inclined pipe (16) for reuse.

[0075] The oil gas is transferred to a subsequent fractionation unit (18) via the main oil gas pipeline (17), and after separation, the obtained dry gas is discharged through the pipeline (19); the obtained liquefied gas is discharged through the pipeline (20) and separated into propylene, propane, and C4 hydrocarbons in the gas separation unit (25), and the propylene, propane, and C4 hydrocarbons are discharged through pipelines (26, 27, and 28), respectively; the obtained gasoline is discharged through the pipeline (21); the light circulating oil fraction having a distillation range of 200 to 250°C is recovered through the pipeline (22) and then recirculated through the pipeline (31) to the middle-upper part of the first reaction zone (8) of the variable diameter fluidized bed reactor (2) along with a spray stream from the pipeline (32); with the sprayed steam from the pipeline (32); The obtained slurry oil is recovered through the pipeline (24) and recirculated to the first reaction zone (8) of the variable diameter fluidized bed reactor (2) for purification (optionally delivered to the first reaction zone (8) along with feed oil from the pipeline (3) through the feed nozzle), to recover the fine powder of the catalyst; the obtained catalytic cracking distillate is delivered to the hydrogenation treatment unit (29) through the pipeline (23), and the hydrogenated distillate obtained after hydrogenation treatment is recovered through the pipeline (30) along with heavy deasphalted oil from the pipeline (10') as a blending component for marine fuel. The distillation range and the method of processing each fraction can be adjusted according to the actual needs of the refinery, for example, gasoline can be split to obtain a light gasoline fraction, and the light gasoline fraction can be recirculated through the pipeline (6) to the second reaction zone of the variable diameter fluidized bed reactor (2) along with a spray stream from the pipeline (5) for purification to increase the yield of propylene.

[0076] In a preferred specific embodiment, the present invention provides the following technical solutions:

[0077] 1. A method for manufacturing more propylene and low-sulfur fuel oil components comprises the following steps:

[0078] (1) A step of obtaining deasphalted oil and deoiled asphalt by contacting a solvent for extraction separation with heavy feedstock oil;

[0079] (2) a step of supplying deasphalted oil and optionally light feedstock oil as catalytic conversion feedstock oil to a catalytic conversion reactor and contacting it with a catalytic conversion catalyst for the reaction to obtain liquefied gas including propylene, gasoline, and catalytic cracking distillate;

[0080] (3) A step of obtaining low-sulfur hydrogenated distillate by hydrodesulfurizing catalytic cracking distillate;

[0081] Here, the low-sulfur hydrogenated distillate and / or deoiled asphalt is used as a fuel oil component.

[0082] 2. In the method of item 1, the heavy feedstock oil is selected from vacuum residue, degraded atmospheric pressure residue, hydrogenated heavy oil, or a mixture of two or more of these.

[0083] 3. In the method of item 1, the solvent is selected from lower alkanes or a mixture of two or more of these, wherein the lower alkanes are selected from the group consisting of propane, butane, and pentane, or a mixture of two or more of these.

[0084] 4. In the method of item 1, step (1) is carried out at an operating temperature of 10-200°C, preferably 20-180°C, an operating pressure of 1.0-15.0 MPa, preferably 2.0-10.0 MPa, and a mass ratio of solvent to feedstock oil of 1-20, preferably 3-10.

[0085] 5. In the method of item 1, the light feedstock oil used in step (2) is selected from petroleum hydrocarbons and / or other mineral oils, the petroleum hydrocarbons are selected from the group consisting of vacuum gas oil, atmospheric gas oil, coker gas oil, high-quality residue, high-quality hydrogenated heavy oil, or a mixture of two or more of these, and the other mineral oils are selected from the group consisting of coal liquefaction oil, tar sand oil, shale oil, or a mixture of two or more of these.

[0086] 6. In the method of item 1, at step (2), the reactor is selected from a group consisting of combinations of two or more reactors of the same type, including riser reactors, constant linear velocity fluidized bed reactors, equal diameter fluidized bed reactors, upward conveyor lines, downward conveyor lines, combinations of two or more of these, or combinations of reactors connected in series and / or parallel, wherein the riser reactor is a conventional equal diameter riser reactor or various types of variable diameter fluidized beds.

[0087] 7. In the method of item 1, the catalytic conversion catalyst used in step (2) comprises, based on the total weight of the catalyst, about 1-50 weight% of zeolite, about 5-99 weight% of inorganic oxide, and about 0-70 weight% of clay, wherein the zeolite is a mesoporous zeolite and optionally a macroporous zeolite, wherein the mesoporous zeolite accounts for 51-100 weight% of the total weight of the zeolite, the mesoporous zeolite has a silica-alumina ratio greater than 50, preferably greater than 80, and the macroporous zeolite accounts for 0-49 weight% of the total weight of the zeolite.

[0088] 8. In the method of item 1, the catalytic conversion condition of step (2) is a reaction temperature of 460-750°C, preferably 480-700°C; 10-100 h -1 , preferably 30-80 h -1It includes a weight-hour space velocity of 4-20, preferably 5-12, of catalyst-converted feedstock oil.

[0089] 9. In the method of item 1, the catalytic cracking distillate obtained in step (2) has an initial boiling point of 200°C or higher and a hydrogen content of 12.0% by weight or less.

[0090] 10. In the method of item 9, the catalytic cracking distillate has an initial boiling point of 250°C or higher and a hydrogen content of 11.0% by weight or less.

[0091] 11. In the method of item 1, the catalyst used in the hydrodesulfurization step (3) is a catalyst comprising a group VIB metal and / or a group VIII metal supported on an alumina and / or amorphous silica-alumina carrier.

[0092] 12. In the method of item 11, the hydrogenation catalyst comprises 0-10 wt% of an additive, 1-40 wt% of one or more Group VIII metals, 1-50 wt% of one or more Group VIB metals, and the remainder being an alumina and / or amorphous silica-alumina carrier, wherein the additive is selected from the group consisting of non-metallic elements such as fluorine, phosphorus, etc., and metallic elements such as titanium, platinum, etc.

[0093] 13. In the method of item 1, the hydrodesulfurization conditions are a reaction pressure of 2.0-24.0 MPa, a reaction temperature of 200-500 ℃, and 50-5000 Nm 3 / m 3 Volume ratio of hydrogen to oil, and 0.1-30.0 h -1 It includes the liquid's space velocity per hour.

[0094] 14. In the method of Item 13, the hydrodesulfurization conditions are a reaction pressure of 3.0-15.0 MPa, a reaction temperature of 300-400 ℃, and 200-2000 Nm 3 / m 3 The volume ratio of hydrogen to oil, and 0.2-10.0 h-1 It includes the liquid's space velocity per hour.

[0095] 15. In the method of item 1, the hydrogenated distillate obtained in step (3) has a sulfur content of 0.1 weight% or less, preferably 0.05 weight% or less.

[0096] Examples

[0097] The present invention will be described in more detail with reference to the following examples, but is not limited thereto.

[0098] The characteristics of the feedstock oil and catalyst used in the following examples and comparative examples are shown in Tables 1 and 2, respectively. The catalyst conversion catalyst used in the comparative example was MMC-1, a catalyst manufactured by the Qilu Branch of Sinopec Catalyst Co., Ltd.

[0099] The hydrogen content of the catalytic cracking distillate obtained in each example was measured using a carbon and hydrogen analysis meter according to the NB / SH / T 0656-2017 standard.

[0100] The catalytic conversion catalyst used in the examples was prepared according to the following.

[0101] 969 g of halloysite (available from China Kaolin clay Co., Ltd., solid content 73%) was slurried in 4300 g of deionized water, 781 g of pseudo-boehmite (available from Shandong Zibo Bauxite Plant, solid content 64%) and 144 ml of hydrochloric acid (concentration 30%, specific gravity 1.56) were added, and the mixture was stirred uniformly. The mixture was left to stand and aged at 60°C for 1 hour, maintaining the pH value at 2-4, and then the mixture was cooled to room temperature. 5000 g of a pre-prepared slurry containing 1600 g of mesoporous shape-selective ZSM-5 zeolite (available from the Qilu Branch of Sinopec Catalyst Co., Ltd.) containing chemical water and a silica-alumina ratio higher than 150 was added, uniformly stirred, and the product was spray-dried and washed to remove free Na+ to obtain a catalyst. The obtained catalyst was aged with 100% steam at 800°C, the aged catalyst was named Catalyst A, and its characteristics are shown in Table 2.

[0102] The hydrogenation desulfurization catalyst B used in the example was prepared as follows:

[0103] After weighing 1000g of pseudo-boehmite manufactured by the ChangLing Branch of Sinopec Catalyst Co., Ltd., 1000ml of an aqueous solution containing 10ml of chemically pure nitric acid was added. The mixture was band-extruded using a twin-screw extruder to form the material, dried at 120°C for 4 hours, and calcined at 800°C for 4 hours to obtain a catalyst support. The support was immersed in 900ml of an aqueous solution containing 120g of ammonium fluoride for 2 hours, dried at 120°C for 3 hours, and calcined at 600°C for 3 hours; after cooling to room temperature, the product was further immersed in 950ml of an aqueous solution containing 133g of ammonium metamolybdate for 3 hours, dried at 120°C for 3 hours, and calcined at 600°C for 3 hours; After cooling the product back to room temperature, the product was finally immersed in 900 ml of an aqueous solution containing 180 g of nickel nitrate and 320 g of ammonium metatungstate for 4 hours, dried at 120°C for 3 hours, and calcined at 600°C for 4 hours to prepare catalyst B.

[0104] [Table 1] Characteristics of feedstock oils used in the examples and comparative examples

[0105]

[0106] [Table 2] Characteristics of catalyst conversion catalysts used in examples and comparative examples

[0107]

[0108] Example 1-a

[0109] This example was carried out according to the process method shown in FIG. 1 using vacuum residue VR-1 as a heavy feedstock oil. The heavy feedstock oil was solvent deasphalted with propane, and the characteristics of the obtained deasphalted oil and de-oiled asphalt are shown in Table 3.

[0110] Using Catalyst A as a catalytic conversion catalyst, a test was performed on a medium-sized catalytic cracking unit comprising a variable-diameter fluidized bed reactor with 100% propane deasphalted oil. The generated oil gas and spent catalyst were separated in a separator, and the oil gas products were separated into propylene, butylene, gasoline, and catalytic cracking distillate (distillation range 250-500°C, hydrogen content 9.4 wt%) depending on the distillation range of the fractions within the fractionation unit. The reaction conditions and product distribution are listed in Table 4.

[0111] The obtained catalytic cracking distillate is sent to a hydrodesulfurization reactor with hydrogen and contacted with hydrodesulfurization catalyst B, at a reaction pressure of 6.0 MPa, a reaction temperature of 350°C, a hydrogen-to-oil volume ratio of 350, and a liquid space velocity of 2.0 h -1 A low-sulfur hydrodistillate was obtained by reacting in [a specific way]. The low-sulfur hydrodistillate was used as a fuel oil component and blended with a second fuel oil component (i.e., degreased asphalt obtained in this example) and a third fuel oil component, i.e., hydrodiesel oil, to obtain a fuel oil product RMG 180 that meets the National Standard GB 17411-2015, Marine Fuel Oils, and its characteristics are shown in Table 5.

[0112] Example 1-b

[0113] This example was carried out as described in Example 1, except that a medium-sized unit including an equal-diameter riser reactor was used instead of a medium-sized unit including a variable-diameter fluidized bed reactor. The generated oil gas and spent catalyst are separated in a separator, and the oil gas products are separated into propylene, butylene, gasoline, and catalytic cracking distillate (distillation range 250-500°C, hydrogen content 9.4 wt%) according to the distillation range of the fractions in the fractionation unit. The reaction conditions and product distribution are listed in Table 4.

[0114] Example 2

[0115] This example was carried out according to the process method shown in FIG. 1 using vacuum residue VR-1 as a heavy feedstock oil. The heavy feedstock oil was solvent deasphalted with butane, and the characteristics of the obtained deasphalted oil and de-oiled asphalt are shown in Table 3.

[0116] Using Catalyst A as a catalytic conversion catalyst, a medium-sized catalytic cracking unit comprising a variable diameter fluidized bed reactor with a mixture of 80% butane deasphalted oil and 20% VGO was tested. The generated oil gas and spent catalyst were separated in a separator, and the oil gas products were separated into propylene, butylene, gasoline, and catalytic cracking distillate (distillation range 250-500°C, hydrogen content 10.1 wt%) according to the distillation range of the fractions within the fractionation unit. The reaction conditions and product distribution are listed in Table 4.

[0117] The obtained catalytic cracking distillate is sent to a hydrodesulfurization reactor with hydrogen and contacted with hydrodesulfurization catalyst B, at a reaction pressure of 7.0 MPa, a reaction temperature of 380°C, a hydrogen-to-oil volume ratio of 500, and a liquid space velocity of 1.5 h -1 Low-sulfur hydrodistillate was obtained by reacting in [a specific way]. The low-sulfur hydrodistillate was used as a fuel oil component and blended with another fuel oil component (i.e., degreased asphalt obtained in this example) to obtain a fuel oil product RMG 380 that meets the National Standard GB 17411-2015, Marine Fuel Oils, and its characteristics are shown in Table 6.

[0118] Example 3

[0119] This example was carried out according to the process method shown in FIG. 1 using vacuum residue VR-1 as a heavy feedstock oil. The heavy feedstock oil was solvent deasphalted with pentane, and the characteristics of the obtained deasphalted oil and de-oiled asphalt are shown in Table 3.

[0120] Using Catalyst A as a catalytic conversion catalyst, tests were performed on a medium-sized catalytic cracking unit comprising a variable diameter fluidized bed reactor with a mixture of 60% pentane deasphalted oil and 40% VGO. The generated oil gas and spent catalyst were separated in a separator, and the oil gas products were separated into propylene, butylene, gasoline, and catalytic cracking distillate (distillation range 250-500°C, hydrogen content 10.4 wt%) depending on the distillation range of the fractions within the fractionation unit. The reaction conditions and product distribution are listed in Table 4.

[0121] The obtained catalytic cracking distillate was sent to a hydrodesulfurization reactor with hydrogen and contacted with hydrodesulfurization catalyst B, at a reaction pressure of 8.0 MPa, a reaction temperature of 310°C, a hydrogen-to-oil volume ratio of 550, and a liquid space velocity of 4.0 h -1 Low-sulfur hydrodistillate was obtained by reacting in [a specific way]. The low-sulfur hydrodistillate was used as a fuel oil component and blended with another fuel oil component (i.e., degreased asphalt obtained in this example) to obtain the fuel oil product RMG 180, which meets the National Standard GB 17411-2015, Marine Fuel Oils, and its characteristics are shown in Table 7.

[0122] Comparative Example 1

[0123] This comparative example was carried out in a medium-sized unit comprising a riser reactor combined with a dense fluidized bed, using VGO as the feedstock oil and catalyst MMC-1 as the catalytic cracking catalyst, according to the conventional deep catalytic cracking process described in CN1004878B. The generated oil gas and spent catalyst were separated in a separator, and the oil gas products were separated into propylene, butylene, gasoline, and catalytic cracking distillate (distillation range 200-350°C, hydrogen content 9.8 wt%) according to the distillation range of the fractions in the fractionation unit. The reaction conditions and product distribution are listed in Table 4.

[0124] [Table 3] Solvent deasphalting conditions and results performed in Examples 1-3

[0125]

[0126] [Table 4] Conditions and product distribution of catalytic conversion reactions performed in Examples 1-3 and Comparative Example 1

[0127]

[0128]

[0129] * Conversion rate = Dry gas yield + Liquefied gas yield + Gasoline yield + Coke yield.

[0130] As can be seen from the results in Table 4, Examples 1-a and 1-b can provide a propylene yield of 5 wt% or more, as well as a fuel oil component yield of about 70 wt% (calculated based on hydrodistilled oil + degreased asphalt for vacuum residue VR-1 used as heavy feedstock oil). Compared to Comparative Example 1, the dry gas yield of Examples 1-a and 1-b is significantly reduced, and the total liquid yield is significantly increased.

[0131] [Table 5] Characteristics of the low-sulfur hydrogenated distillate and fuel oil products obtained in Example 1-a

[0132]

[0133] [Table 6] Characteristics of low-sulfur hydrogenated distillate and fuel oil products obtained in Example 2

[0134]

[0135] [Table 7] Characteristics of the low-sulfur hydrogenated distillate and fuel oil products obtained in Example 3

[0136]

[0137] Example 4

[0138] This example was carried out according to the process method shown in FIG. 2 using hydrogenated heavy oil as a heavy feedstock oil. The heavy feedstock oil was solvent deasphalted with butane, and the characteristics of the light deasphalted oil, heavy deasphalted oil, and deasphalted asphalt are shown in Table 8.

[0139] Using Catalyst A as a catalytic conversion catalyst, a test was performed on a medium-sized catalytic cracking unit comprising a variable-diameter fluidized bed reactor with 100% light butane deasphalted oil. The generated oil gas and spent catalyst were separated in a separator, and the oil gas products were separated into propylene, butylene, gasoline, and catalytic cracking distillate (distillation range 250-500°C, hydrogen content 10.4 wt%) depending on the distillation range of the fractions within the fractionation unit. The reaction conditions and product distribution are listed in Table 9.

[0140] The obtained catalytic cracking distillate was sent to a hydrodesulfurization reactor with hydrogen and contacted with hydrodesulfurization catalyst B, at a reaction pressure of 9.0 MPa, a reaction temperature of 330°C, a hydrogen-to-oil volume ratio of 650, and a liquid space velocity of 8.0 h -1 Low-sulfur hydrodistillate was obtained by reacting in [a specific way]. The low-sulfur hydrodistillate was used as a fuel oil component and blended with another fuel oil component, “vacuum residue VR-2,” to obtain the fuel oil product RMG 180, which meets the National Standard GB 17411-2015, Marine Fuel Oils, and its characteristics are shown in Table 10.

[0141] Example 5

[0142] This example was carried out according to the process method shown in FIG. 2 using hydrogenated heavy oil as a heavy feedstock oil. The heavy feedstock oil was solvent deasphalted with propane, and the characteristics of the light deasphalted oil, heavy deasphalted oil, and deasphalted asphalt are shown in Table 8.

[0143] Using Catalyst A as a catalytic conversion catalyst, a test was performed on a medium-sized catalytic cracking unit comprising a variable-diameter fluidized bed reactor with 100% light propane deasphalted oil. The generated oil gas and spent catalyst were separated in a separator, and the oil gas products were separated into propylene, butylene, gasoline, and catalytic cracking distillate (distillation range 250-500°C, hydrogen content 10.5 wt%) depending on the distillation range of the fractions within the fractionation unit. The reaction conditions and product distribution are listed in Table 9.

[0144] The obtained catalytic cracking distillate is sent to a hydrodesulfurization reactor with hydrogen and contacted with hydrodesulfurization catalyst B, at a reaction pressure of 6.0 MPa, a reaction temperature of 350°C, a hydrogen-to-oil volume ratio of 350, and a liquid space velocity of 4.0 h -1 A low-sulfur hydrodistillate was obtained by reacting in [a specific way]. The low-sulfur hydrodistillate was used as a fuel oil component and blended with a second fuel oil component (i.e., the heavy deasphalted oil obtained in this example) and a third fuel oil component (vacuum residue VR-3) to obtain a fuel oil product RMG 380 that meets the National Standard GB 17411-2015, Marine Fuel Oils, and its characteristics are shown in Table 11.

[0145] [Table 8] Solvent deasphalting conditions and results performed in Examples 4-5

[0146]

[0147] [Table 9] Catalytic conversion reaction conditions and product distribution performed in Examples 4-5

[0148]

[0149] * Conversion rate = Dry gas yield + Liquefied gas yield + Gasoline yield + Coke yield.

[0150] [Table 10] Characteristics of the low-sulfur hydrogenated distillate and fuel oil products obtained in Example 4

[0151]

[0152] [Table 11] Characteristics of the low-sulfur hydrogenated distillate and fuel oil products obtained in Example 5

[0153]

[0154] As can be seen from the results shown in the table above, the method of the present invention can provide a high-value-added product, namely propylene, while producing a certain amount of fuel oil components.

[0155] Although preferred embodiments of the present invention have been described in detail above, the present invention is not limited to these embodiments. Various modifications are possible without departing from the essence of the present invention, and such modifications are also within the scope of the present invention.

[0156] It should be noted that the various technical features described in the above embodiments can be combined in any appropriate manner without contradiction. For the sake of brevity, various possible combinations are not separately described in this invention, but such combinations are also included within the scope of this invention.

[0157] Furthermore, various embodiments of the present invention may be combined in any manner without departing from the spirit of the present invention, and such combinations should be considered part of the disclosure of the present invention.

Claims

Claim 1 A method for producing propylene and low-sulfur marine fuel oil components comprising the following steps: (1) contacting a heavy feedstock oil with a solvent for extraction separation to obtain light deasphalted oil, heavy deasphalted oil, and deoiled asphalt; (2) contacting the light deasphalted oil with a catalytic conversion catalyst for reaction in a catalytic conversion reactor in the absence of hydrogen to obtain a reaction product containing propylene, and adjusting the conversion rate of the catalytic conversion process to 50 to 70%; (3) separating the reaction product from step (2) to obtain catalytic cracking distillate oil, wherein the catalytic cracking distillate oil has an initial boiling point of 200°C or higher and a final boiling point of 550°C or lower, and a hydrogen content of 12.0 wt% or less; (4) hydrodesulfurizing the catalytic cracking distillate oil to obtain low-sulfur hydrodistillate oil; and (5) a step of obtaining marine fuel oil by blending low-sulfur hydrodistillate with one or more fuel oil components selected from heavy deasphalted oil, vacuum residue and degassed asphalt, wherein the low-sulfur hydrodistillate and / or degassed asphalt are suitable for use as fuel oil components, and the catalytic conversion catalyst used in step (2) comprises 1-50 wt% zeolite, 5-99 wt% inorganic oxide, and 0-70 wt% clay based on the total weight of the catalyst, and the reaction conditions of step (2) are a reaction temperature of 460-750°C or 480-700°C, and 10-100 h -1 , or 30-100 h -1 A method of preparation comprising a space velocity per unit weight of 1-10 seconds, or a reaction time of 2-8 seconds; and a weight ratio of catalyst to oil of 4-20, or 5-12. Claim 2 A method for manufacturing, wherein the zeolite comprises 51-100 wt% of mesoporous zeolite and 0-49 wt% of macroporous zeolite based on the total weight of the zeolite in the catalytic conversion catalyst, the mesoporous zeolite has a silica-alumina ratio of greater than 10, greater than 50, or greater than 100, or the mesoporous zeolite is selected from the group consisting of ZSM-type zeolite and ZRP zeolite; and the macroporous zeolite is a Y-type zeolite. Claim 3 A method of manufacturing, wherein, in claim 1, step (2) is performed such that the resulting reaction product has a propylene / propane mass ratio of 4 or more, or 6 or more, or 8 or more; and / or an isobutene / isobutane mass ratio of 1 or more, or 1.5 or more, or 1.8 or more. Claim 4 A manufacturing method according to claim 1, wherein step (2) is performed such that the yield of catalytic cracking distillate in the generated reaction product is 15% or more, or 20% or more, or 30% or more and 50% or less. Claim 5 A method of manufacturing according to claim 1, wherein the heavy feedstock oil is selected from the group consisting of vacuum residue, hydrogenated heavy oil, or any mixture thereof. Claim 6 A method of manufacturing according to claim 1, wherein the solvent is selected from the group consisting of propane, butane, pentane, or any mixture thereof. Claim 7 A method for manufacturing, wherein the extraction separation conditions of step (1) include a temperature of 10-200 ℃ or 20-180 ℃, an operating pressure of 1.0-15.0 MPa or 2.0-10.0 MPa, and a mass ratio of solvent to feedstock oil of 1-20 or 3-10. Claim 8 A method for manufacturing, wherein the light feedstock oil used in step (2) is selected from the group consisting of petroleum hydrocarbons, other mineral oils, or mixtures thereof, the petroleum hydrocarbons are selected from the group consisting of vacuum gas oil, atmospheric gas oil, coker gas oil, residues with a hydrogen content of 11.2 weight% or more, hydrogenated heavy oil with a hydrogen content of 11.2 weight% or more, or any mixture thereof, and the other mineral oils are selected from the group consisting of coal liquefaction oil, tar sand oil, shale oil, or any mixture thereof. Claim 9 A method of manufacturing according to claim 1, wherein the catalytic conversion reactor used in step (2) is a fluidized bed reactor comprising a single fluidized bed reactor or a complex reactor comprising a plurality of fluidized bed reactors connected in series or parallel, or, a uniform diameter riser reactor or various types of variable diameter fluidized bed reactors. Claim 10 A method of manufacturing, wherein the catalytic cracking distillate of step (3) has an initial boiling point of 250°C or higher, a final boiling point of 520°C or lower, or 500°C or lower, and a hydrogen content of 11.0% by weight or lower. Claim 11 A method of manufacturing according to claim 1, wherein a catalyst comprising a group VIB metal and / or a group VIII metal supported on an alumina and / or amorphous silica-alumina carrier is used in the hydrodesulfurization step (4). Claim 12 A method of manufacturing, wherein the catalyst used in the hydrodesulfurization step (4) comprises 0-10 wt% of an additive, 1-40 wt% of at least one Group VIII metal (calculated as a metal oxide), and 1-50 wt% of at least one Group VIB metal (calculated as a metal oxide), and the remainder being a carrier selected from alumina and amorphous silica-alumina, wherein the additive comprises an element selected from the group consisting of fluorine, phosphorus, titanium, platinum, or a combination thereof. Claim 13 In claim 1, the conditions of the hydrodesulfurization step (4) are a reaction pressure of 2.0-24.0 MPa, a reaction temperature of 200-500 ℃, and 50-5000 Nm 3 / m 3 Volume ratio of hydrogen to oil, and 0.1-30.0 h -1 It includes a liquid hourly space velocity, or the conditions of the hydrodesulfurization step (4) include a reaction pressure of 3.0-15.0 MPa, a reaction temperature of 300-400 ℃, and 200-2000 Nm 3 / m 3 The volume ratio of hydrogen to oil, and 0.2-10.0 h -1 A manufacturing method comprising the liquid hourly space velocity of Claim 14 A method of manufacturing, wherein the hydrogenated distillate obtained in step (4) has a sulfur content of 0.1% by weight or less, or 0.05% by weight or less. Claim 15 A method of manufacturing, wherein step (2) is performed such that the resulting reaction product has a propylene / propane mass ratio of 4 or more, or 6 or more, or 8 or more; and / or an isobutene / isobutane mass ratio of 1 or more, or 1.5 or more, or 1.8 or more. Claim 16 A manufacturing method according to paragraph 2, wherein step (2) is performed such that the yield of catalytic cracking distillate in the generated reaction product is 15% or more, or 20% or more, or 30% or more and 50% or less. Claim 17 In claim 15, step (2) is performed such that the yield of catalytic cracking distillate in the generated reaction product is 15% or more, or 20% or more, or 30% or more and 50% or less.

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